MX2010009562A - Process. - Google Patents

Abstract

A process for activating a reduced manganese copper catalyst comprising treating the catalyst at a temperature of more than 300°C to about 400°C with hydrogen.

Description

PROCESS TO ACTIVATE A COPPER CATALYST WITH REDUCED MANGANESE DESCRIPTIVE MEMORY The present invention relates to a process for activating a catalyst, and a catalyst activated according to this process. Also, the present invention relates to the production of ethers, optionally with the co-production of diols and / or lactones by reaction of an organic feedstock in the presence of hydrogen. In general, the reaction is carried out by hydrogenation and / or dehydration. The organic feedstock is selected from dicarboxylic acids and / or anhydrides, monoesters of dicarboxylic acids and / or anhydrides, diesters of dicarboxylic acids and / or anhydrides, lactones, a mixture thereof, or a mixture of two or more of the same. In particular, it relates to the production of C4 to C12 ethers, optionally with the co-production of the corresponding diols and / or lactones by the reaction of di- (C1 to C4 alkyl) esters of C4 to C12 dicarboxylic acids and / or anhydrides in the presence of hydrogen. More particularly, it relates to the production of cyclic ethers.

More particularly, the present invention relates to a process for the co-production of C4 compounds, more specifically tetrahydrofuran, butane-1,4-diol and / or β-butyrolactone from a material hydrocarbon prime comprising a dialkyl maleate by vapor phase reaction in a stream enriched in hydrogen. In a particularly preferred embodiment of the present invention, it relates to a process for the production of at least 20% tetra-id breakage not with the co-production of butane-1,4-diol and / or β-butyrolactone. In the most preferred embodiment, it is related to the production of tetrahydrofuran where any residual butane-1,4-diol and / or residual β-butyrolactone is recycled and converted to additional tetrahydrofuran.

It is known that diols are produced by hydrogenation of dialkyl esters of dicarboxylic acids and / or anhydrides, lactones, and mixtures thereof with a minor amount, typically not more than about 10% w / w and preferably not more than 1% w / w p, of a monoester of dicarboxylic acid and / or anhydride. Commercial plants have been constructed which produce butane-1,4-diol as the main product with small amounts, typically up to about 10 mole%, tetrahydrofuran and up to about 15 mole% β-butyrolactone by hydrogenation of an ester of dialkyl maleic acid and / or anhydride, such as dimethyl maleate or diethyl maleate, which may contain minor amounts of dialkyl fumarate and / or dialkyl succinate. Dimethyl succinate or diethyl succinate were also suggested as starting materials suitable for hydrogenation to produce butane-1,4-diol, tetrahydrofuran and β-butyrolactone. These succinates can be formed in any suitable way and can be from source biotechnology For additional information regarding the operation of these plants, reference may be made to, for example, US-A-4584419, US-A-4751334, WO-A-86/03189, WO-A-88/00937. , US-A-4767869, US-A-4945173, US-A-4919765, US-A-5254758, US-A-53I0954 and WO-A-91/01960, the disclosure of which is incorporated herein by reference .

While many plant operators seek to maximize the performance of butane-1,4-diol and minimize the performance of co-products, tetrahydrofuran and? -butyrolactone, these co-products are, in themselves, commercially valuable chemicals. Tetrahydrofuran, in general, is recovered because it is an important monomer for making elastomeric fibers and is also an important solvent and, consequently, a commercially important chemical. The? -butyrolactone can be recovered but, as the market for this product is reduced, it is often recycled to the hydrogenation step for conversion to additional butane-1,4-diol and the tetrahydrofuran co-product.

The dialkyl maleates which are used as raw material in said hydrogenation processes can be produced by any suitable means. The hydrogenation of dialkyl maleates to obtain butane-1,4-diol is discussed in detail in US-A-4584419, US-A-4751334 and WO-A-88 / Q0937, which are incorporated herein by reference. reference.

A significant part of butane-1,4-diol produced by methods conventionally it is subsequently converted to tetrahydrofuran. This conversion step has substantial cost implications both in the investment and in the operation of the plant required for the conversion and, as the importance of tetrahydrofuran increases along with its use in derived applications, it is desirable to provide a process for the production of tetrahydrofuran without the need for this expensive downstream processing. Downstream processing of conventional methods includes recovering butane-1,4-diol, reacting it to form tetrahydrofuran, and then refining the tetrahydrofuran product.

Processes typically conventions produce up to about 10 mole% tetrahydrofuran.

It is therefore desirable to provide a process for the production of a higher mole% tetrahydrofuran without the need for downstream processing.

A proposal to increase the amount of tetrahydrofuran produced is described in WO 03/00644. In this process, the feed material is fed to a vaporization zone where it is vaporized by and within the cycle gas. The resulting stream is fed into a first reaction zone comprising a catalyst in which hydrogenation and dehydration is carried out. A stream of intermediate product is recovered, and passes to a second vaporization zone where additional feed material is added. The resulting current crosses an area of additional reaction where the hydrogenation and dehydration is carried out.

The process is preferably carried out in the presence of a copper catalyst activated with reduced manganese.

While the above process is successful in increasing the proportion of the tetrahydrofuran produced, there is a problem associated with the rigidity of the catalyst in the face of minor changes in operating conditions.

In extreme cases, as the conditions are altered, the sites where tetrahydrofuran is produced on the catalyst cease to function and can be considered as having been removed. Proposals for catalysts that are more resistant to changes in operating conditions have been suggested, however, they generally do not offer the required level of conversion and / or selectivity.

Another problem with the conventional catalysts appears when the process is operated at higher temperatures such as those that may be required to increase the formation of tetrahydrofuran by over 90%, since there is an increase in the formation of by-products.

It has now been discovered that if a reduced manganese copper catalyst is activated under a stream of hydrogen at temperatures of about 300 ° C to about 400 ° C, a catalyst is obtained which offers advantages in processes for the production of ethers.

Therefore, according to a first aspect of the present invention, a process for activating a copper catalyst with Reduced manganese comprising treating the catalyst at a temperature of more than 300 ° C to about 400 ° C with hydrogen. In a preferred embodiment, the temperature can be from more than 300 ° C to about 330 ° C.

By "reduced manganese" is meant that the catalyst comprises less than 0.1% by weight manganese, more preferably 0.05% by weight or less, such as 0.03% by weight manganese. The reduced manganese copper catalyst is more preferably a copper catalyst with zero manganese.

It has been found that if the catalyst is used in a process for the production of an ether by reaction of a corresponding organic feedstock selected from dicarboxylic acids and / or anhydrides, monoesters of dicarboxylic acids and / or anhydrides, diesters of dicarboxylic acids and / or anhydrides, lactones, and mixtures of two or more thereof in the presence of hydrogen, then higher temperatures, such as from about 215 ° C to about 230 ° C are conventionally used to improve ester conversion; improved selectivity is noted with respect to the desired ether. For example, the conversion of essentially 100% of the desired ether can be achieved. Under these conditions, a high byproduct formation, as indicated by the butanol content, would be expected. However, surprisingly, when the catalyst has been activated in accordance with the present invention, the butanol content is lower than that obtained in similar conversions, but a a temperature 20 ° C lower. It was also found that the activated catalyst has acceptable longevity without degradation of ether production over time. Since the reaction can be operated at higher temperatures, the impact of water present in the feed is also reduced.

A preferred activation process comprises the steps of: (i) supplying a stream comprising a maximum of 0.5% hydrogen to the catalyst at room temperature; (ii) increasing the temperature to a temperature in excess of 300 ° C for a period of 10 to 20 hours; Y (iii) increase the hydrogen content of the stream until it is 100%.

The particular advantages are reached where the current supplied in step (i) starts at 0.1% hydrogen and, subsequently, increases in steps to 0.5% in a period of 5 to 10 hours, more preferably approximately 7 hours. In a more preferred embodiment, as the hydrogen content in the stream supplied in step (i) increases, the ambient temperature increases to a temperature in the range of about 100 ° C to about 160 ° C.

It may be desirable to monitor and carefully adjust the hydrogen inlet and outlet content to control the exotherm.

According to the second aspect of the present invention, an activated catalyst according to the first aspect above is provided. described.

According to the third aspect of the present invention, a process is provided for the production of an ether by reaction of a corresponding organic feedstock selected from dicarboxylic acids and / or anhydrides, monoesters of dicarboxylic acids and / or anhydrides, diesters of dicarboxylic acids and / or anhydrides, lactones, and mixtures of two or more thereof in the presence of hydrogen, wherein the reaction is carried out in the presence of a catalyst of the second aspect of the present invention, or an activated catalyst in accordance with with the process of the first aspect described above.

In a preferred example of the third aspect of the present invention, the process comprises the steps of: (a) supplying a stream comprising the organic feedstock to a first vaporization zone and contacting said feedstock with cycle gas comprising hydrogen in such a way that at least a portion of the feedstock is vaporized by and within the cycle gas; (b) supplying the cycle gas and the vaporized feedstock to a first reaction zone comprising a catalyst, wherein the reaction zone operates under reaction conditions to allow hydrogenation and dehydration to take place, (c) recovering from the first reaction zone a stream of intermediate product comprising feed material without react, cycle gas, desired product (s), and any co-products and by-products; (d) supplying the intermediate product stream to a second vaporization zone and bringing it into contact with additional feedstock such that said additional feedstock is vaporized by and into the intermediate product stream; (e) supplying the product of step (d) to a subsequent reaction zone comprising catalyst and operating under reaction conditions to allow hydrogenation to be carried out and, if necessary, dehydration; Y (f) recovering from the subsequent reaction zone a product stream comprising the ether, wherein at least one of the catalyst of step (b) and (e) comprises the catalyst of the second aspect of the present invention, described above.

In one embodiment, the catalysts used in steps (b) and (e) may be different. When the catalysts used in step (b) and (e) are different, the catalyst used in step (b) can be an acid-tolerant catalyst such as an activated copper chromite catalyst and that for step (e) can be the catalyst of the second aspect of the present invention. A suitable catalyst for use in step (b) is the catalyst available from Davy Process Technology Ltd. as PG85 / 1.

Although the previous preferred process has been described with Particular reference to two reaction zones, in one embodiment of the present invention, the process may include more than two reaction zones. When there are more than two reaction zones, the corresponding vaporization zones can be located between the adjacent reaction zones. The vaporization in these subsequent zones can be carried out directly in the intermediate product stream from the previous reaction zone or, if a complementary cycle gas stream is required which can comprise one or more of a new organic feed, it can be Include refining and hydrogen recycling material. The feed and / or hydrogen recycle material, if present, can be heated.

When these intermediate reaction zones are present, they may include the same catalysts of step (b) or (e) or, in an alternative embodiment, a different catalyst may be used. In one embodiment, the catalyst may be one that is effective to hydrogenate the ester to diols and lactones such as manganese activated copper catalyst. A suitable manganese activated copper catalyst is one available from Davy Process Technology Ltd. as DRD 92 / 89A. This catalyst exhibits superior conversion of a dialkyl ester under typical operating conditions.

The catalyst used in the reaction zones can be a single catalyst or a mixture of catalysts. In a particularly preferred process, the catalyst of the first reaction zone may include a noble metal and / or copper-containing catalysts. Thus, The catalyst of the first hydrogenation zone can be or include one or more of a palladium catalyst, a chromium reduced copper catalyst or a reduced copper containing catalyst. The same or a different catalyst can be used in the subsequent reaction zone or any additional reaction zone.

Examples of copper-containing catalysts include zinc oxide / reduced copper oxide catalysts, reduced manganese activated copper catalysts, reduced copper chromite catalysts and activated reduced copper chromite catalysts.

The catalytic active species can be at least partially supported on a support material selected from chromium, zinc oxide, alumina, silica, silica-alumina, silicon carbide, zirconium, titanium, carbon, or a mixture of two or more of the themselves, for example, a mixture of chromium and carbon.

In one embodiment, a bed comprising a variety of catalysts may be used provided that at least one catalyst in a bed is the catalyst of the second mentioned aspect of the present invention. In an example, the bed may include a catalyst that is tolerant to the acid content of the waste feed, one which is suitable for promoting the hydrogenation of the ester and the activated copper catalyst which does not include manganese is used which promotes selectivity to the desired ether. Catalyst beds comprising more than one type of catalyst may comprise discrete layers of catalyst in the bed, such that the Different types are separated or the different types of catalysts can be mixed.

In cases where the ester feed contains acid components, a protective bed of a suitable catalyst may be present to hydrogenate the acid and protect the catalyst of the second aspect of the present invention.

In the ether production reaction of the present invention, the conversion of the acid, anhydride and / or the lactone or ester to the diol is a hydrogenolysis or ester hydrogenation and the reaction of the diol to the ether is a dehydration reaction.

Without intending to conform to any theory, it is believed that the preferred process allows the amount of product produced to be increased as light boiling ether (higher vapor pressure) instead of diol, such that the dew point of the The reactor is moved below the operating temperature in such a way that the additional feed material can be vaporized in the stream until the current reaches saturation. This is in contrast to conventional processes for the production of diols in which the inlet and outlet of the reactor are close to the condensation point of the vapor. The additional feed material vaporized by the process of the present invention can then be converted into the product in the second reaction zone.

The cycle gas usually contains a high concentration of hydrogen gas but can also include other gases that include hydrocarbons, carbon oxides, methane, nitrogen. Also, when the cycle gas includes recycled gases from the downstream, condensables may be present which include ether of the product, Ci to C4 alkanol, water, co-products and by-products.

In a particularly preferred aspect of the third embodiment of the present invention, the ether is a cyclic ether. More preferably, the cyclic ether is tetrahydrofuran. In the latter case, the organic feedstock is preferably dialkyl maleate. Co-products that may be present to a greater or lesser degree in this embodiment, or that may be absent, include butane-1,4-diol and β-butyrolactone. This reaction is illustrated in Scheme 1. In this example, alkanol is methanol and the intermediate material is partially hydrogenated dimethyl succinate.

The by-products may include the alkanol used in the esterification of the acid or anhydride, for example methanol, undesired material formed in the side reactions, for example butanol, water emitted in the dihydration of the diol to the ether and intermediate material, for example succinate of Dimethyl together with other light or heavy materials formed in the process.

COOMe COOMe + 2H2 • 2MeOH H SCHEME 1 The by-products can be separated from the ether in a refining zone and can be further purified, if necessary. Similarly, the co-products can be separated from the ether in the refining zone and can be further purified, if necessary.

However, in one embodiment, one or more of the co-products and / or by-products are recycled to the vaporization zone where they will be vaporized. In an alternative embodiment, one or more of the co-products and / or by-products are recycled to the second vaporization zone where they will be vaporized in the intermediate product stream leaving the first reaction zone.

Therefore, in the preferred embodiment, any succinate of The dialkyl present as a by-product can be recycled to the first vaporization zone and thus to the first reaction zone to improve the total selectivity of the reaction to the desired tetrahydrofuran and the butane-1,4-diol and / or co-products. -butyrolactone.

The cycle gas recovered from the subsequent reaction zone is preferably compressed, recycled and mixed with hydrogen formed before being heated and recycled to the vaporization zone.

Any liquid alkanol stream separated from the product can be recycled to an upstream esterification zone.

The feed material to the vaporization zone, or each of them, may be, or may include, one or more recycle streams. Fresh organic feed streams and refined recycle streams can be vaporized together or they can be vaporized in separate parts of the vaporization zone, or each of them. This is particularly advantageous since it minimizes the transesterification risk between the ester and the diol.

In one embodiment, all cycle gas and organic feed fed to the first vaporization zone (step a) are supplied to the first reaction zone (step b) where the remaining organic feed and the refining recycles are vaporized ( step d) in the intermediate product stream recovered from the first reaction zone (step c) to form the intermediate feed stream which is fed to the subsequent reaction zone (step d).

In a second alternative embodiment, the gaseous stream from the first vaporizer (step a) can be divided, wherein a larger part, preferably from about 70% to about 80%, is supplied to the first reaction zone (step b) and a smaller part, preferably from about 20% to about 30%, bypassing (by-passing) the first reaction zone and being fed to the subsequent vaporization zone, preferably a part of the subsequent vaporization zone (step d), where it is further heated in such a manner that the additional organic feed material can be vaporized in the cycle gas before forming a secondary hot feed stream. When the smaller part is fed to a part of the subsequent vaporization zone, the intermediate product stream recovered from the first reaction zone (step c) is fed to a second part of the subsequent vaporization zone (step d) in the which refined recycles are fed. The two streams from the two separate parts of the subsequent vaporization zone are then mixed to form the intermediate feed stream which is fed to the subsequent reaction zone (step e).

An advantage of this preferred embodiment is that the additional liquid organic feed, which may be or may include an ester, is separated from the liquid refining recycles containing diols and / or lactones, and is mixed only with them in the vapor phase . This minimizes the contact time and, therefore, the potential for transesterification and the progressive chain length growth.

The feed material fed to the vaporization zone, or each of them, can be, or can include, completely one or more of the recycle streams.

The cycle gas usually contains a high concentration of hydrogen gas but can also include other gases that include hydrocarbons, carbon oxides, methane, nitrogen. Further, when the cycle gas includes recycled gases from the downstream, condensables may also be present which include ether of the product, Ci to C4 alkanol, water, co-products and by-products.

The feed material fed to the vaporization zone may be, or may include, one or more recycle streams. The new organic feed and refining recycle streams can be vaporized together or vaporized in separate portions of the vaporization zone. This is particularly advantageous since it minimizes the transesterification risk between the ester and the diol.

The organic feed material is preferably selected from mono C1 to C4 alkyl esters of C4 to C12 dicarboxylic acids and / or anhydrides, di- (C1 to C4 alkyl) esters of C to C12 dicarboxylic acids and / or anhydrides, lactones of hydrocarboxylic acids C to C12, and mixtures of two or more thereof.

For example, organic feed material can be selecting from mono alkyl Ci to C4 esters of C4 dicarboxylic acids and / or anhydrides, di- (Ci to C4 alkyl) esters of C4 dicarboxylic acids and / or anhydrides,? -butyrolactone, and mixtures of two or more thereof. A particularly preferred organic feedstock can be selected from monomethyl maleate, monomethyl fumarate, monomethyl succinate, dimethyl maleate, dimethyl fumarate, dimethyl succinate, β-butyrolactone, recycle β-butyrolactone and / or butane-1. , 4-diol, and mixtures of two or more thereof. Alternatively, the organic feedstock can be selected from monoethyl maleate, monoethyl fumarate, monoethyl succinate, diethyl maleate, diethyl fumarate, diethyl succinate, β-butyrolactone, recycle and / or butanoy-butyrolactone. -1, 4-diol, and mixtures of two or more thereof.

In one embodiment, the organic feed material fed to the vaporization zone is contained in an organic solvent. When the organic solvent is present, the vaporization zone operates in such a way that the organic feed material is essentially separated from the organic solvent by stripping the cycle gas.

Suitable organic solvents include: di (alkyl Ci to C4) esters of alkyl dicarboxylic acids containing up to 13 carbon atoms; mono- and di- (alkyl Ci to C4) esters of maleic acid, fumaric acid, succinic acid, and mixtures thereof; (alkyl Ci to C4) esters of naphthalene monocarboxylic acids, tri- (alkyl Ci to C4) esters of aromatic tricarboxylic acids; di- (alkyl Ci to C) esters of isophthalic acid; alkyl phthalates; Y dimethyl sebecato.

The stream to the first reaction zone preferably has a cycle gas ratio containing hydrogen: vaporized feed in the range of about 50: 1 to about 1,000: 1. Typically, the stream to the first reaction zone is at a temperature of from about 100 ° C to about 300 ° C, more preferably from about 150 ° C to about 250 ° C. Any suitable pressure can be used, but the feed pressure to the first reaction zone is typically from about 50 psia (about 344.74 kPa) to about 2,000 psia (about 13,789 kPa). In one embodiment, the pressures are in the range of about 450 psia (approximately 3.102.64 kPa) to 1,000 psia (approximately 6,894.76 kPa). The feed to the first reaction zone is preferably supplied to the first reaction zone at a rate corresponding to a space velocity per hour of liquid of about 0.05 h "1 to about 5.0 h'1.

If necessary, the pressure and / or temperature can be adjusted in any convenient manner between the reaction zones. The temperature can be adjusted by any suitable means including the use of an exchanger or heat exchangers.

The hydrogen gas formed used in the process of the present invention can be obtained by any conventional manner. Preferably, it contains at least about 50% by volume up to about 99.99% by volume or more, for example from about 80 to about 99.9% by volume, of hydrogen. It may also contain one or more inert gases, such as nitrogen or methane. Conveniently, the hydrogen gas formed is produced by oscillating pressure absorption such that the molecular weight of the cycle gas is minimized, thus reducing the energy required for compression and circulation of the cycle gas.

Typically, the hydrogenatable material contains from about 0.01 to about 1.0% w / w or more, for example up to about 10% w / w, but usually not more than about 2.0% w / w, of acidic material.

The charge of the catalyst in the first reaction zone is preferably large enough to reduce the content of acidic material to less than about 0.005% w / w in the passage of the vapor mixture therethrough.

The amount of catalyst used in each reaction zone can be the same or different. The charge of the catalyst in the first reaction zone may constitute from about 10% to about 70%, more usually from about 20% to about 50%, of the total volume of the catalyst in the reaction zones. Similarly, the catalyst of the subsequent reaction zone is typically in the range of about 70% to about 10%, more usually from about 20% to about 50%, of the total volume of the catalyst of the reaction zones.

The selected catalyst preferably converts the ester, preferably the dialkyl maleate, to the desired ether, preferably a cyclic ether, more preferably tetrahydrofuran, at a selectivity of about 20% to about 90% or more, more preferably, about 70% or more. The selectivity approaching 00% is particularly desirable.

The product stream from the final reaction zone is preferably fed, having been preferably condensed, to a refining zone where the desired ether, preferably tetrahydrofuran, is separated as a product. Any co-product, such as butane-1,4-diol and / or β-butyrolactone, which may be present may be separated or recycled to the reaction system. When there is more than one co-product, one or more can be separated and recovered, and the remainder can be recycled.

In an embodiment where 100% conversion to ether is desired, for example tetrahydrofuran, all co-products, for example butane-, 4-diol and / or β-butyrolactone, are recycled.

Any alkanol derived from the organic feed, which is typically a C-to C4 alkanol and water in the crude product stream, preferably condenses and separates in the raffinate. The alkanol is recycled in conventional manner to an esterification reactor in which the organic feedstock is formed, if present. The refining system it may include means, if necessary, to separate the water from the alkanol. The refining system usually includes means to separate other by-products that can be recycled. An example of a by-product that can be recycled is, for example, any intermediate material. Alternatively, some or all of any of the by-products formed can be rejected as effluent. An example of a by-product that can be recycled is any mono-ol produced.

In cases where wetting of the catalyst can cause the catalyst to deteriorate, it may be desirable to feed the reaction mixture to the reactor above the dew point. This can be achieved by passing a suitable excess cycle cycle gas through the vaporizer or by adding extra cycle gas flow after the vaporizer, or by adding extra heat to the reaction mixture before feeding the reaction zone. However, if the wetting of the catalyst is not detrimental to the operation of the catalyst, entrained liquid may be present. However, the reaction will still essentially be a vapor phase reaction.

Other details of a suitable process can be found in WO 03/00644, which is incorporated herein by reference.

The invention is further described with reference to the following examples.

EXAMPLE 1 In order to activate the copper catalyst, 500 ml of DRD 92/89"D" catalyst available from Davy Process Technology Limited was placed in the reaction zone. The gas velocity was established at 1,250 NLPH. The pressure was set at 50 psig. The following procedure was subsequently performed: 1. The concentration of hydrogen was increased to 0. 1% at room temperature. The inlet temperature was then increased to 120 ° C in 3 hours. The inlet / outlet hydrogen composition was monitored once the temperatures were above 100 ° C and the hydrogen input was maintained at 0.1%.

In steps 2 to 6, the exotherm was prevented from exceeding 10 ° C by reducing the hydrogen input composition, if necessary, and kept under these conditions until the exotherm was reduced. 2. The temperature increased by 10 ° C until it reached 160 ° C 3. After holding at 160 ° C for 1 hour, the hydrogen composition of the inlet gas was increased to 0.2% in 1 hour and kept under these conditions for 2 hours, 4. The hydrogen content of the inlet composition was increased by 0.3% and maintained for 2 hours. 5. The hydrogen content of the input composition it increased by 0.4% and remained for 2 hours. 6. The hydrogen content of the inlet composition was increased by 0.5% and until the hydrogen content of the inlet composition equaled that of the hydrogen outlet. 7. The hydrogen inlet composition of 0.5% was maintained, and the temperature was increased to 340 ° C in 12 hours. It was ensured that the exotherm did not exceed 10 ° C and was maintained until the hydrogen input equaled the hydrogen output. 8. The temperature was maintained at 340 ° C for 1 hour. 9. The hydrogen input was increased by 1% in a period of 1 hour and was maintained until the hydrogen input equaled the hydrogen output. The exotherm was kept below 10 ° C reducing the concentration of hydrogen, if necessary. 10. The concentration of hydrogen at the inlet was then increased to 5% in 4 hours. 1 1. The hydrogen content at the inlet was then increased to 10% and maintained until the hydrogen input equaled the hydrogen output. The exotherm remained below 10 ° C. 12. The concentration of hydrogen was then increased to 100%, ensuring that the exotherm did not exceed 10%. 13. The pressure was increased to the operating pressure of the reaction. The catalyst was then cooled to the required temperature and allowed to stand under hydrogen for 4 hours before feeding.

EXAMPLES 2 TO 5 The DRD 92/89 D catalyst obtained from Davy Process Technology Ltd. and activated as described above was used in a process for the production of tetrahydrofuran from a feed comprising maleic anhydride. The reaction conditions are detailed in Table 1, and the results in Table 2.

TABLE 1 Note: LHSV is Space Speed by Liquid Hour (stands for Liquid Hourly Spacial Velocity).

TABLE 2 Selectivities % mol Example 2 Example 3 Example 4 Example 5 Tetrahydrofuran 83.21 2.69 92.78 97.43 n-butanol 0.15 0.15 0.25 0.52 G-Butyrolactone 1 1 .79 13.38 4.07 0.13 1. 4-butanediol 1 .63 1 .85 0.70 0.00 Conversion. % in moles 67.72 67.48 91 .59 99.28 EXAMPLE 6 The DRD 92/89 D catalyst obtained from Davy Process Technology Ltd. and activated as described in advance was used in combination with the DRD 92/89 A catalyst from Davy Process Technology. The combined catalyst was used in a process for the production of tetrahydrofuran from a feed comprising maleic anhydride. The reaction conditions are detailed in Table 3, and the results in Table 4.

COMPARATIVE EXAMPLE 7 The DRD 92/89 D Catalyst obtained from Davy Process Technology Ltd. was activated in accordance with conventional methods and was used in combination with the DRD 92/89 A catalyst from Davy Process Technology. The combined catalyst was used in a process for the production of tetrahydrofuran from a feed comprising maleic anhydride. The reaction conditions are detailed in Table 3, and the results in Table 4.

Despite the fact that in Comparative Example 7 double the Type D catalyst is used as used in Example 6, the selectivity with respect to tetrahydrofuran is substantially improved when the catalyst has been subjected to the process of activating the I presented invention.

TABLE 3 Note: LHSV is Space Speed per Liquid Hour (stands for the English Liquid Hourí and Spacial Velocity).

TABLE 4 EXAMPLE 8 The DRD 92/89 D catalyst obtained from Davy Process Technology Ltd. was activated as described in advance and was used in a process for the production of tetrahydrofuran from a feed which comprises maleic anhydride. The reaction conditions are detailed in Table 5, and the results in Table 6.

COMPARATIVE EXAMPLES 9 AND 10 The DRD 92/89 D catalyst obtained from Davy Process Technology Ltd. was activated in accordance with conventional low temperature processes. The catalyst was used in a process for the production of tetrahydrofuran from a feed comprising maleic anhydride. The reaction conditions are detailed in Table 5, and the results in Table 6.

TABLE 5 Note: LHSV is Space Speed per Liquid Hour (stands for the English Liquid Hourí and Spacial Velocity).

TABLE 6 Selectivities, mol% Example Example 9 Comparative Example 8 Comparative 10 Tetrahydrofuran 83.21 45.1 69.8 1. 4-butanediol 0.15 29.1 19.2 G-Butyrolactone 11.79 21.2 7.2 Conversion% 67.7 50.6 60.1

Claims (14)

NOVELTY OF THE INVENTION CLAIMS

1. A process for activating a reduced manganese copper catalyst comprising treating the catalyst at a temperature of more than 300 ° C to about 400 ° C with hydrogen, comprising the steps of: (i) supplying a stream comprising a maximum of 0.5% hydrogen to the catalyst at room temperature; (ii) increasing the temperature to a temperature in excess of 300 ° C for a period of 10 to 20 hours; and (iii) increasing the hydrogen content of the stream until it is 100%; wherein the current supplied in step (i) is initiated at 0.1% hydrogen and then increased stepwise to 0.5% over a period of 5 to 10 hours, and where as the hydrogen content of the stream increases Supplied in step (i), the ambient temperature is increased to a temperature in the range of about 100 ° C to about 160 ° C.

2. process in accordance with the claim further characterized in that the temperature is from more than 300 ° C to approximately 330 ° C.

3. The process according to claim 1 or 2, further characterized in that the reduced manganese copper catalyst is a copper catalyst with zero manganese.

4. An activated catalyst according to the process of any of claims 1 to 3.

5. A process for the production of an ether by reaction of a corresponding organic feedstock, selected from dicarboxylic acids and / or anhydrides, monoesters of dicarboxylic acids and / or anhydrides, diesters of dicarboxylic acids and / or anhydrides, lactones, and mixtures of two or more thereof in the presence of hydrogen, wherein the reaction is carried out in the presence of the catalyst of claim 4 or is activated according to the process of any of claims 1 to 3.

6. The process according to claim 5, further characterized in that it comprises the steps of: (a) supplying a stream comprising the organic feed material to a first vaporization zone and contacting said feed with cycle gas comprising hydrogen from such that at least a part of the feedstock is vaporized by and within the cycle gas; (b) supplying the cycle gas and the vaporized feedstock to a first reaction zone comprising a catalyst, wherein the reaction zone operates under reaction conditions to allow hydrogenation and dehydration, (c) recovering starting from the first reaction zone a stream of intermediate product comprising unreacted feedstock, cycle gas, desired product (s), and any other co-product and by-product; (d) supply the intermediate product stream to a second vaporization zone and contacting it with additional feed material such that said additional feed material is vaporized by and into the intermediate product stream; (e) supplying the product from step (d) to a subsequent reaction zone comprising catalyst and operating under reaction conditions to allow hydrogenation to be carried out and, if necessary, dehydration, and (f) recover from the subsequent reaction zone a product stream comprising the ether; wherein at least one of the catalyst of step (b) and (e) comprises the catalyst of claim 4 or is activated according to the process of any of claims 1 to 3.

7. The process according to claim 6, further characterized in that the catalyst used in step (b) comprises an acid-tolerant catalyst and that of step (e) comprises the catalyst of claim 4 or is activated in accordance with the process of any of claims 1 to 3.

8. The process according to claim 7, further characterized in that the catalyst used in step (b) comprises an activated chromite and copper catalyst.

9. The process according to any of claims 6 to 11, further characterized in that the process includes one or more additional subsequent reaction zones located in series between the first and the subsequent subsequent reaction zone and wherein the final subsequent reaction zone, or each of them, is preceded by a zone of vaporization in which the additional feed, new feed or recycle and recycle are vaporized by or within the intermediate product stream from the previous reaction zone.

10. The process according to claim 9, further characterized in that the catalyst in the intermediate reaction zone is a copper catalyst activated with manganese.

11. The process according to any of claims 6 to 10, further characterized in that the feedstock is selected from mono alkyl Ci to C esters of C4 to C12 dicarboxylic acids and / or anhydrides, di- (C1 to C4 alkyl) esters of C4 to C12 dicarboxylic acids and / or anhydrides, C to C12 hydrocarboxylic acid lactones, and mixtures of two or more thereof.

12. The process according to claim 11, further characterized in that the feedstock is selected from monomethyl maleate, monomethyl fumarate, monomethyl succinate, dimethyl maleate, dimethyl fumarate, dimethyl succinate, y-butyrolactone, β-butyrolactone. of recycle and / or butane-1,4-diol and mixtures of two or more thereof.

13. The process according to any of claims 6 to 12, further characterized in that the ether is a cyclic ether.

14. The process according to claim 13, further characterized in that the ether is tetrahydrofuran.